Process for producing styrenic/alkenylnitrile copolymers

ABSTRACT

Disclosed is a process and apparatus for the continuous mass polymerization of styrenic and alkenylnitrile monomers. The process comprises the steps of continuously introducing a feed comprising a predetermined ratio of styrenic and alkenylnitrile monomers into a reaction vessel to produce a reaction mixture; subjecting the reaction mixture to conditions of temperature and pressure under which said monomers copolymerize to produce a styrenic/alkenylnitrile copolymer; subjecting the reaction mixture to agitation; continuously withdrawing styrenic/alkenylnitrile copolymer from the reaction vessel; and by withdrawing vapor phase from the reaction vessel in a controlled amount by varying the amount of cooled surface available for condensing the vapors, condensing the vapors by contact with the cooled surface and returning the condensate to the reaction vessel. The process and apparatus also include a novel means for devolatilizing the product copolymer.

RELATED APPLICATIONS

This application is a continuation-in-part of parent application Ser.No. 423,488 filed on Sept. 24, 1982, now abandoned, which is herebyincorporated by reference in its entirety.

FIELD OF THE INVENTION

The present invention relates to an improved system for the continuousmass polymerization of copolymers. More specifically, the presentinvention relates to a process for the continuous polymerization ofstyrenic-alkenylnitrile copolymers (SAN.)

BACKGROUND OF THE INVENTION

The copolymerization of vinyl monomers, particularly styrene andacrylonitrile is well known, as are the processes for their production.However, there are still many inherent problems which have either beenunsolved, or in many cases, solved unsatisfactorily.

In a copolymerization reaction, precise control is desirable, if notessential, to maintain the quality of the product. For example, ifmonomer X is to be copolymerized with monomer Y, a wide range ofcopolymers are possible. The X monomer could comprise from 1 to 99% ofthe copolymer. Incident to this is a wide variation in the physical andchemical properties of the product copolymers. Accordingly, precisecontrol is required such that the final product consists essentially ofcopolymers having a substantially uniform X:Y ratio. The closer theuniformity, or the narrower the range of variance, the better theproduct. Even in those instances where a wide range is tolerated, it iswell known that a narrower range of variance results in a betterproduct.

In a typical continuous mass copolymerization process, the monomers feedis introduced into a reaction vessel. The copolymerization is achievedunder elevated temperature and pressure conditions. The reaction vesselis the first major area where control is essential. A uniformtemperature needs to be maintained to produce a uniform product. Theresultant fluid contains the produced copolymer and part of each of themonomers. Withdrawal of the fluids from the reaction vessel presents thesecond problem area. Since the copolymerization reaction in the reactionvessels is not complete, some monomers are still prsent in the withdrawnfluids. If uncontrolled, some copolymerization would continue to takeplace to produce a copolymer having a different monomers ratio due tothe different temperatures and pressure conditions. Thus, control ofthis continued copolymerization is essential to keep it to a minimum.Further, if any continued copolymerization occurs, control over theconditions should be exercised such that the produced second stagecopolymer has substantially the same monomers ratio.

After the withdrawn fluids are cooled, the separation of the copolymersfrom the monomers and other diluents takes place. This is normallyachieved by applying heat to vaporize the undesired constituents. Thus,a third problem area is present. Heat application would causecopolymerization or further reaction of the constituents which leads tothe contamination of the final product. Thus, this third polymerizationshould be kept to a minimum.

For example, in the case of SAN polymerization, it is known that precisecontrol is essential to the production of a product having acceptableproperties. If the acrylonitrile content of copolymer in a singleproduct varies by more than about 4%, the different copolymers becomeincompatible with one another, resulting in an unacceptable, hazyproduct. Because styrene and acrylonitrile monomers polymerize atdifferent rates, careful control is needed not only during thepolymerization stage, but also during the subsequent purificationstages. Especially in these latter processing stages, it is essential tomaintain uniformity of the product, because even small amounts ofcopolymer product having a high acrylonitrile content can cause a yellowdiscoloration of the entire product, due to cyclization of adjacentpendant acrylonitrile groups upon heating of the copolymer, e.g., evenduring subsequent thermoforming steps.

This is an important factor in the postpolymerization treatment of theproduct in a SAN polymerization process, e.g., the removal of residualmonomer from the polymer, known as devolatilizing the polymer. Once thepolymer/monomer mixture leaves the reactor there is great risk ofproducing high-acrylo-nitrile-containing polymer, due to the unevenrates of polymerization for the two different monomer species and thehigh temperatures which are utilized for devolatilization. Inconventional processes, devolatilization is typically carried out withthin film devolatilizing equipment, such as the so-called "Film Truder",which evaporates the liquid monomer very rapidly to minimize furtherpolymerization. This thin film equipment, however, is relativelyexpensive and requires an inordinate amount of maintenance, e.g., atleast once daily, because of its many moving parts and the extensiveseals characteristic thereof. This maintenance requires, therefore, thatthe polymerization line be shut down or that some measure be taken tohold material upstream during maintenance. The result is an unevenquality of product and added expense for equipment and operation.

SUMMARY OF THE INVENTION

It is therefore an object of the present invention to provide animproved continuous mass polymerization apparatus.

Another object of the invention resides in providing an improved processand apparatus for maintaining a uniform temperature in the reactorutilized for the copolymerization of vinyl monomers.

In accomplishing the foregoing objects, there has been provided inaccordance with one aspect of the present invention, a process forcontrolling reaction temperature in a continuous vinyl monomerspolymerization reaction comprising introducing a feed comprising vinylmonomers into a reaction vessel to produce a reaction mixture. Thereaction mixture is subjected to conditions of temperature and pressureunder which the monomers copolymerize. The reaction mixture in thereaction vessel is cooled by withdrawing vapor phase from the reactionvessel to a condenser capable of retaining condensed liquids wherein thevapor phase is cooled by contact with a cooled condensing surface toproduce a condensed liquid and wherein the rate of withdrawal of vaporphase from the reaction vessel is controlled by controlling the cooledsurface area available for contact with the withdrawn vapor phase.

In accordance with another aspect of the present invention, a processfor the continuous mass polymerization of styrenic and alkenylnitrilemonomers to produce a styrenic/alkenylnitrile copolymer, to produce acondensed liquid and wherein the rate of withdrawal of vapor phase fromthe reaction vessel is controlled by controlling the cooled surface areaavailable for contact with the withdrawn vapor phase.

In accordance with another aspect of the present invention, a processfor the continuous mass polymerization of styrenic and alkenylnitrilemonomers to produce a styrenic/alkenylnitrile copolymer is provided. Afeed comprising a predetermined ratio of styrenic and alkenylnitrilemonomers is continuously introduced into a reaction vessel to produce areaction mixture. The reaction mixture containing styrenic andalkenylnitrile monomers is subjected to conditions of temperature andpressure under which said monomers copolymerize to produce a liquidcontaining styrenic/alkenylnitrile copolymer and styrenic andalkenylnitrile monomers. The reaction mixture is subjected to agitationsufficient to maintain a substantially uniform composition distributionand a substantially uniform temperature distribution throughout thereaction mixture. At least part of the liquid containingstyrenic/alkenylnitrile copolymer and the respective monomers iscontinuously withdrawn from the reaction vessel. The reaction mixture inthe reaction vessel is cooled by withdrawing vapor phase containingvaporized styrenic and alkenylnitrile monomer from the reaction vesselto a condenser. The vaporized styrenic and alkenylnitrile monomerswithdrawn from the reaction vessel are condensed by contact with acooled condensing surface to produce a condensed monomer-containingliquid. At least a part of the condensed monomer-containing liquid isreturned to the reaction vessel. The cooling step further comprisescontrolling the amount of vaporized monomers withdrawn from the reactionvessel by varying the amount of cooled surface area available forcontact with the withdrawn vaporized monomers in response to thetemperature in the reaction vessel, whereby the amount of vaporizedmonomers condensing is controlled.

Further objects, features and advantages of the present invention willbecome apparent from the detailed description of preferred embodimentswhich follows, when considered together with the attached figures anddrawings.

BRIEF DESCRIPTION OF THE DRAWINGS

In the drawings:

FIG. 1 is a schematic view of a suitable apparatus for carrying out theprocess according to the present invention;

FIG. 2 is a more detailed schematic view of an alternative apparatus forcarrying out the controlled temperature polymerization according to theinvention; and

FIG. 3 is a schematic view illustrating a preferred apparatus forcondensing and recycling recovered monomers according to the invention.

DETAILED DESCRIPTION OF THE PREFERRED EMBODIMENTS

While the following description and examples are specific to thecopolymerization of styrenic and alkenylnitrile monomers, it should beclear that the present invention is applicable to copolymerizationprocesses in general particularly where control over copolymercomposition is desirable.

The present invention is directed to an improved continuous masspolymerization process for making copolymers which utilizes a startingmonomer composition comprising, at least one monoalkenyl aromatic(styrenic) compound and at least one alkenylnitrile compound. As usedherein, the term "mass polymerization" does not exclude the presence ofdiluents in the process.

The alkenylnitrile compounds are characterized by the general formula:##STR1## wherein R is selected from the group consisting of hydrogen andalkyl radicals containing from 1 through 4 carbon atoms each.

The monoalkenyl aromatic (styrenic) compounds are characterized by thegeneral formula: ##STR2## wherein Ar is selected from the groupconsisting of a phenyl radical, an alkaryl radical of 6 through 9 carbonatoms, a monochlorophenyl radical, a dichlorophenyl radical, amonobromophenyl radical, and a dibromophenyl radical, and

X is selected from the group consisting of hydrogen and an alkyl radicalcontaining less than three carbon atoms.

The preferred alkenylnitrile is acrylonitrile, and the preferredstyrenic is styrene itself.

The invention provides a process and apparatus which enable theefficient production of a wide range of SAN copolymer products under awide range of operating conditions. For example, according to theinvention, SAN copolymers can be prepared having a proportion ofcopolymerized acrylonitrile ranging from as little as about 10% byweight up to a maximum of about 60% by weight. A wide range of molecularweights is also possible, e.g., varying from polymers having a melt flowindex (ASTM 1238, condition I) as low as about 1 gram/10 min. up to ashigh as 80 gram/10 min. This wide range of values can be achievedwithout the use of internal lubricants in the polymer compositions. Thepolymers are characterized as having a Vicat softening point (ASTM 1525Rate A) falling within the range of 220°-250° F.

SAN copolymers of excellent purity and quality are producible accordingto the invention. These compositions have a residual acrylonitrilemonomer content of less than about 100 ppm, a residual styrene monomercontent of less than about 1000 ppm and a residual diluent content ofless than 1000 ppm.

Yet, the process of the invention is also capable of achieving very highrates of reaction and conversion levels. It is possible to achieve areaction rate as high as 40%/hour for the SAN copolymers, and conversionlevels of between 30 and 90%, typically up to about 60%.

A. Reactor

With reference to FIG. 1 of the drawings, the system utilizes apolymerization reactor vessel 1. The configuration may vary widely, forexample, from a configuration which is essentially horizontal to onewhich is essentially vertical. The preferred design comprises avertical-type reactor, preferably a vertical cylinder, inasmuch as ithas been found that a vertical reactor enables production of a widerrange of SAN products. In the past, horizontal type reactors have beenpreferred for SAN, because of the difficulty in achieving uniformtemperature and concentration control from the top to the bottom of avertical reactor. See, e.g., U.S. Pat. No. 3,813,369. Many types anddesigns of reactors are known in the art for the polymerization of SAN.It is possible to employ any of these known reactors within the contextof the present invention.

The reactor may be jacketed to aid in heat control, especially duringstart-up. Hot oil is used as the heat exchange medium and it iscirculated through external circuit 6. In the instant invention,however, it does not normally play a significant role in temperaturecontrol during steady state operation. After start-up, the temperatureof the jacket is typically kept at the desired reaction temperature, inorder to minimize temperature gradients.

In operation, a mixture of styrene and acrylonitrile monomers is chargedinto the reactor through line 7, being initially supplied from a supplysource 4 and passing through preheater 8. The ratio, for example, ofstyrene to acrylonitrile in the reaction mixture is selected accordingto the desired monomeric composition of the copolymer. Typically,compositions comprise about 60-70% by weight styrene and about 30-40% byweight acrylonitrile. The reaction mixture typically also includes aninert diluent, for example, from about 2 to 50% by weight and typicallyfrom about 15 to 50% by weight of the total reaction mixture; about 25%is preferred. The percentage may be as low as 2%; however, below thislevel, plugging can become a problem due to the lack of the diluent. Thediluent is preferably selected from the following compounds:ethylbenzene, butylbenzene, benzene, toluene, xylene (all isomers) andcumene; ethylbenzene is preferred. The reactor is usually operated at atemperature of between about 115° and 180° C. and at a pressure ofbetween about 40 and 80 psia.

Referring now to FIG. 2, there is illustrated one preferred reactordesign for the present invention. The reactor 1 includes a combinationof measures to aid in accomplishing a uniform temperature andcomposition therein. Inside the reactor is a rotary mixer 3. The mixerhas at least one blade or paddle 4 which rotates in a horizontal plane.The mixer may be of any conventional design which will assure lateraluniformity within the reaction medium. The mixer is driven by a motor 2driving a rotating shaft 5 extending through the top of the reactor anddown into the reaction mixture. The mixer is typically operated atspeeds between about 20 and 60 rpm.

Lateral uniformity of the reaction mixture is achieved by the rotarymixer 3. Vertical uniformity of the reaction mixture is accomplishedaccording to the invention by withdrawing reaction mixture from thebottom of the reactor and pumping the mixture via pump 10 through anexternal recycle line 9 in an external loop and then back into the topof the reactor 1. The reaction mixture as it passes through the loop ispreferably homogenized in mixing device 11, preferably a static mixer ofthe type referred to as an "interfacial surface generator". An"interfacial surface generator" is an inline motionless mixer, whosemixing mechanism is generally unrelated to the throughput when thethroughput is flowing in the region of streamline flow. Such mixers maybe considered as layering mixers wherein the flowing stream is dividedand two component parts reshaped and joined together in such a way thatthe interface between the original elements of the stream issubstantially increased. Such mixers are well known in the art and someof these mixers and their mode of operation are described in thefollowing U.S. Pat. Nos. 3,051,542; 3,051,453; 3,195,865; 3,206,170;3,239,197; 3,286,992; 3,328,003; 3,358,749; 3,373,534; 3,394,924;3,404,869; 3,406,947; and 3,506,244. This in-line static mixer may be ofany standard design. Suitable mixers are sold by various suppliers underthe trade names of Kenics, Koch, and Lightning.

Fresh monomer is introduced into the recycle loop through line 13,upstream of the static mixer 11, as needed to maintain a constantmonomeric ratio within the reactor. As will be discussed in more detailbelow, condensed monomer vapor is also returned to the polymerizationvessel by introduction upstream of the static mixing device 11.

The use of the external pumping loop 9 and static mixer 11 incombination with the rotary mixer 3 provides surprising uniformity oftemperature and composition throughout the reaction mixture, even in thecase of a vertical reactor. The temperature variation from top to bottomin the reactor can be controlled to within ±1° C. when the reactionmixture is circulated as little as 1.2 times per hour through theexternal loop. Similarly, the monomeric composition of the reactionmixture is maintained uniform to within 1%. A 15,000 ton per year plantutilizing the instant design requires only about 200 HP for the rotarymixer 3 and a pump of approximately 60 HP for the external loop.

B. Evaporative Temperature Control

Under steady-state operating conditions, substantially all of thecooling for the reactor 1 is provided by an external monomer condenser15. Hot monomer vapors are withdrawn from the reactor 1 through line 17and condensed in condenser 15. The condenser 15 may be of eitherhorizontal or vertical design, with a vertical design illustrated inFIG. 1. Within the condenser are a plurality of tubes cooled by acooling fluid, such as water. The monomer vapor may be passed on thetube side or on the shell side; however, in a vertical design it ispreferably passed on the tube side. Upon contact with the cold tubes,the hot vapor from the reactor is condensed into liquid monomer whichcollects inside of condenser 15. Also, an additional quantity of theliquid is preferably collected in collector 19. The cooled liquid incollector 19, which is principally monomer, is passed through line 20 tothree-way valve 21, from which a portion of the condensed monomer isconducted via line 22 back to monomer feed line 13, upstream of staticmixer 11, whereupon the condensed monomer is reintroduced intopolymerization reactor 1.

Control over this evaporative cooling system is provided as follows: Atemperature controller 25 reads directly the temperature in reactor 1and operates in cascade with level controller 24 associated withcondenser 15. The amount of cooling provided to the reactor is afunction of the level of condensed monomer in condenser 15, i.e., thelower the fluid level, the more area of cooling surface is exposed tothe monomer vapors. Thus, as more cooling capacity is needed for thereactor, the liquid level in condenser 15 is lowered, and vice-versa.Adjustment of the liquid level in condenser 15 can be accomplished in anumber of ways, for example, by providing a feedback signal to three-wayvalve arrangement 21. Other means of varying the amount of cooledsurface area may be utilized. For example, the flow of cooling mediumthrough the cooling tubes may be partially halted. Additionally, part ofthe cooling tubes may be mechanically withdrawn from the condenser.Other processes will be evident to those skilled in the art as long asit is understood that the available cooling surface is changed tocontrol the amount of vapor withdrawn from the reaction vessel.

The evaporative cooling arrangement according to the invention actuallyprovides a two-fold mechanism to achieve rapid response to temperaturechanges within the polymerization reactor. As noted above, the coolingcapacity of condenser 15 increases directly as a result of an increasein temperature in the reactor. This is accomplished by lowering thelevel of condensed monomer within the condenser, and as a result, thevolume of condensed monomer displaced from the condenser is injectedback into the reactor, thereby producing a secondary and immediatecooling effect. Furthermore, by maintaining a constant supply ofcondensed monomer in collector 19, the system of the invention alsooffers the potential for instantaneously injecting the entire quantityof condensed monomer collected therein into the reactor, in the eventthat a significant instantaneous cooling load is demanded.

Prior art processes utilizing an external condenser for temperaturecontrol in mass polymerization processes, such as mass SAN processes,characteristically suffer from the serious problems of plugging andvapor lock in line 17 between the reactor and condenser, for the reasonthat the cooling load is typically controlled by a control valve in line17 to control the flow of vapor into the condenser, i.e., this valve wasused as the principal temperature regulator. Line 17 has a tendency tobecome plugged at this valve within a short period of time, e.g., in aslittle as one or two days of operation. With this plugging comes theloss of cooling capability for the polymerization reactor. It wasdetermined that the plugging at the valve results from the entropychange there due to the slight pressure drop. Another of the problemswhich plagued prior art systems using external vapor condenser coolingwas vapor lock in the vapor line 17. This was determined to be caused byaccumulation of non-condensibles in the system, e.g., air.

These problems are solved in part according to the present invention bythe elimination of the valve in line 17. This was especially true forthe plugging problem of the valve, inasmuch as the removal of the valveeliminated the source of the problem. Also, the inlet to the condenser15 should be situated above the reactor to eliminate vapor traps.

As mentioned, part of the cause of vapor lock in such a system was foundto be the presence of oxygen and other non-condensibles. To eliminatenon-condensibles, the condenser has a pressure control valve 31 on thetop thereof. This pressure controller utilizes a trap 30 fornon-condensibles and controls the pressure at just above the vaporpressure of the condensed liquid. Therefore, if any non-condensiblesbuild up in the trap 30, the pressure will increase slightly and will bevented automatically by the pressure control valve 31.

The operation of a preferred means for eliminating non-condensibles isas follows: The pressure control valve 31 is positioned at the top ofcondenser 15 above a trap 30 for non-condensibles, so that any buildupof non-condensibles occurs in trap 30. If the pressure control valve isset at a release pressure just slightly above the vapor pressure of thecondensed liquid, the non-condensibles are automatically vented.

The non-condensibles can also be vented in an alternative arrangement.Since these stagnant non-condensibles have a lower temperature than thehot vapor, then the interface between the non-condensibles and the hotvapor can be sensed by a temperature transducer. As this interfacedescends past the first of a pair of temperature transducers indicatinga buildup of non-condensibles, the change in temperature is sensed andthe non-condensibles are vented. This embodiment can be integrated withthe first embodiment, and in such a combined system the pair oftemperature transducers may operate as an override for the pressurecontroller.

As a further means to protect against condensation, the vapor line 17 isinsulated to assure that there is no cooling in the line.

Another of the problems found in prior art evaporative cooling systemsis plugging of the condenser. While vapor phase polymerization ispractically non-existent, hot liquid styrene and acrylonitrile monomerspolymerize readily. Therefore, the plugging of the tubes of thecondenser occurs because the superheated vapor condenses to very hotliquid on the tubes. This liquid can readily polymerize. The designaccording to the present invention has eliminated this problem and, atthe same time, permits even more precise temperature control than waspossible in the prior art. This is accomplished by recirculation of aportion of the mixture of cooled monomer and diluent to the condenserfrom collector 19. The condensed monomer mixture passes from 19 by line20 through three-way valve 21. This valve splits the monomer flow into astream 22, recirculating into reactor 1, and a second stream 23 flowinginto the condenser 15. This recirculation through line 23 solves theproblem of condenser polymerization. The condensate is injected into thecondenser 15 at the point of entry of the superheated vapor stream fromthe reactor. It is injected in the form of a curtain of spray, throughwhich the vapor passes. This cool spray acts as a quencher for the hotvapors and reduces the heat of the liquid monomers on the cooling tubes,further cooling the liquid monomer condensed on the tubes and reducingthe residence time of condensed monomer on the cooling tubes.

By use of the quench spray, polymerization within the condenser isvirtually eliminated. In a vertical heat exchanger, the entering hotvapor stream passes through the cooling spray. The amount of cooledmonomer passing through line 23 is controlled such that it is greaterthan the amount of monomer condensing from vapor on the cooling tubes.The cooled monomer recycled into the condenser thus forms a film of coolmonomer running down the tubes continually. Advantageously, the coolmonomer is introduced with 180° coverage. In a horizontal condenser, therecycled condensate is advantageously introduced via a pipe running theentire length of the top of the condenser. This pipe would have aplurality of holes or nozzles along its length to enable it to wash theentire length of the cooling tubes within the condenser.

FIG. 2 depicts an alternative embodiment of the reactor and temperaturecontrol system wherein a horizontal condenser is employed. Reactor 1embodies an external recirculation loop 9 including static mixer 11.Condenser 15 is horizontally oriented and includes a plurality of tubes14 through which a cooling medium flows, for example, cooling waterentering through line 16 and exiting through line 18. Monomer-containingvapors exiting from reactor 1 through line 17 enter condenser 15 and arecondensed upon contact with the outside surfaces of tubes 14. Thesurface area of tubes 14 available for contact with themonomer-containing vapors is regulated by the level of condensed liquidretained in condenser 15. This liquid level is controlled by levelcontroller 24 in response to the temperature inside reactor 1 asmonitored by temperature controller 25. Control is through adjustment ofthree-way valve 21, which controls the amount of condensed liquidreturned to reactor 1 through line 22 and the amount recycled to thecondenser 15 through line 23.

Condensed liquid recycled to condenser 15 is fed into a pipe 27 whichruns near the top of condenser 15 and along its entire length. Pipe 27contains a plurality of outlet apertures by means of which the condensedliquid is caused to be sprayed across all of the tubes 14, to preventthe buildup of polymer thereon. Furthermore, the monomer-containingvapors are caused to pass through a curtain of spray as they entercondenser 15 through line 17, whereby the vapors are quenched.

A non-condensibles trap 30 and pressure control valve 31 are providedfor continuously purging the system of non-condensibles, in particular,of oxygen. The alternative system described above can also be used here.

The mass polymerization system as thus far described provides one othervery significant advantage over the prior art. In prior art SANproduction systems, it was necessary that the reactor be made ofstainless steel. A reactor of less expensive carbon steel would quicklycorrode, causing black specks in the polymer. Additionally, carbon steelwas believed to cause popcorn SAN polymer which would plug transferlines and equipment. This problem has been solved by the designaccording to the invention as well, allowing very substantial costsavings. In the system of the invention, as was mentioned above, thenon-condensibles are continuously purged from the reactor. Thiseliminates corrosion due to the presence of oxygen within the reactor.Additionally, it is also a feature of the invention to maintain theamount of water within the system at or below its equilibriumconcentration, i.e., about 5-6%. Thus, any liquid water within thereaction mixture is prevented from contacting the reactor walls bymaintaining the water in solution. Excess water is removed in thecondenser system via water draw 32, which is of conventional design andwhich is located in a trap at the bottom of collector 19. Waterseparates from the relatively cool monomer mixture in collector 19,because the solubility of water in acrylonitrile decreases asacrylonitrile is cooled. Thus, by removing water from the monomermixture at this point, the water concentration within the reactor iskept at a sufficiently low level to ensure that it is completelysolubilized. This removal of water and non-condensibles providessignificant cost advantages over prior art systems by allowing the useof carbon steel reactors. The popcorn polymer problem of the prior artis also eliminated.

The system of the present invention allows the making of SAN copolymersusing a reaction temperature of from 115° to about 180° C., preferably130° to 155° C., also the system operates within a pressure range offrom 40 to 80 psia.

C. Polymer Work-up

A portion of the reaction mixture in reactor 1 is continually withdrawnthrough line 33 for further processing. With a SAN polymerizationprocess, it is necessary to remove the residual volatiles (styrene,acrylonitrile and ethylbenzene) and it is preferable to recycle themback into the reactor. The devolatilized polymer is then pelletized orotherwise processed for commercial use after leaving thedevolatilization stage via line 34.

As noted above, the disagreeable yellow color is produced in the SANpolymer when acrylonitrile groups polymerize in series and then cyclizeupon exposure of the polymer to elevated temperatures. Underpolymerization conditions in the reactor, styrene polymerizes morerapidly. Thus, there is an excess of acrylonitrile monomer in themonomer-polymer mixture which has been withdrawn from the reactor forfurther processing. This increases the likelihood of adjacentpolymerization of acrylonitrile in the copolymer chains, changing themonomeric ratio in the final polymer, and leading to yellowing in thepolymer product. As also noted above, the presence of SAN polymershaving differing monomer ratios causes clouding of the product. Thisproblem is addressed by the prior art by removing the volatiles from thepolymer as quickly and at as low a temperature as possible. Mostcommercial installations use a complex and very expensive thin filmevaporator marketed under the trademark of FILM TRUDER. The FILM TRUDERheats and devolatilizes a thin film of polymer very quickly, in lessthan 1 or 2 minutes. The drawback of the FILM TRUDER is that it has acomplex system of mechanical seals and rotating equipment requiringdaily maintenance and accompanying shutdown or bypassing of theequipment. The system provided in accordance with the present inventioneliminates the disadvantages of prior art processes for an equipmentcost of approximately one-fourth that of the FILM TRUDER.

The reactor 1 is operated at super-atmospheric pressure. A polymersolution is withdrawn from the reactor through line 33 at a temperatureof from about 115° C. to about 180° C., preferably about 140° C. Thepolymer mixture is passed into a first devolatilizer unit 35 withoutpreheating. Devolatilizer 35 operates at a pressure above atmosphericpressure, but at a lower pressure than reactor 1. The pressure drop fromline 33 into devolatilizer 35 causes a portion of the acrylonitrile andstyrene monomers in the hot polymer solution to flash off. Sinceacrylonitrile has a higher vapor pressure than styrene, the amount ofacrylonitrile removed may be controlled by controlling the pressure ofdevolatilizer 35. The pressure in devolatilizer 35 is controlled suchthat acrylonitrile monomer is flashed off in sufficient amounts toprovide a monomer ratio in the solution which will produce, uponpolymerization, a SAN copolymer having substantially the same ratio ofstyrene to acrylonitrile monomer units as the SAN copolymer exiting fromthe reactor. Normally, this would involve restoring the monomer ratio tothe same ratio as that initially fed into the polymerization reactor.With the monomer balance thus restored, small amounts of additionalpolymerization in the subsequent processing steps will not disturb thecopolymer ratio (or homogeneity) in the finished polymer. Devolatilizer35 may be of any design suitable for controllably flashing offacrylonitrile. In one embodiment, the devolatilizer is a simple flashpot or one-stage distillation apparatus. The devolatilizer 35 mayalternatively comprise a vessel having a distribution nozzle to injectthe hot polymer solution into the interior thereof, with provision atthe top for removing vapors and provision at the bottom for removingliquid. The temperature entering devolatilizer 35 is preferably betweenabout 115° and 190° C., and the pressure between about 25-30 psia for aSAN process.

The acrylonitrile monomer-rich overhead product is removed from thedevolatilizer 35 through line 37 for recycling into the reactor via acondenser unit, which will be described more fully below. Thepolymer/monomer solution is removed from the first devolatilizer 35 vialine 39 and is transferred into preheater 41. Preheater 41 is a heatexchanger capable of heating the polymer solution from a temperature ofabout 70° C. (to which the temperature can drop in the firstdevolatilizer 35) up to a temperature of from about 190° to 260° C. Theheated polymer mixture then moves from preheater 41 through line 43 intoa second devolatilizer 45, which operates preferably within this lastmentioned temperature range. Even though the solution now has amonomeric composition which is similar to that being fed to the reactor,the polymerization reaction must not be allowed to go to completion, ora SAN copolymer having a molecular weight lower than that of the productremoved from the reactor would form, causing a decrease in the qualityof the product. Therefore, it is essential that the residence time inpreheater 41 be as low as possible, that there be a low heat history inthe preheater and also that the volatiles be removed as quickly aspossible in the subsequent devolatilizer 45. Toward this end, line 43 ismade of relatively large diameter, sufficient that the absolute pressurefrom devolatilizer 45, which is typically between about 0.2 and 0.3psia, is approximately equivalent in preheater 41. Furthermore, thetubes 42 themselves in the preheater heat exchanger are madesufficiently large that the polymer/monomer solution therein experiencesthe vacuum from the second devolatilizer 45.

Running preheater 41 under a partial vacuum produces surprising andadvantageous results. In the first portion of the tubes 42 in preheater41, the monomer in the polymer solution is heated to boilingtemperature. The boiling temperature is quite low, because of the lowpressure in the heat exchanger. The pressure in the heat exchangerpermits boiling at temperatures as low as about 150° C. Thepolymer/monomer solution now becomes a two-phase fluid, advantageously afoam, and is rapidly heated in the remainder of the heat exchanger tothe temperature necessary for good devolatilization, approximately 230°C. It has surprisingly been found that the two-phase liquid has a heattransfer coefficient of about 6 BTU/hr °F. ft², which is generally aboutthree times higher than that of the liquid polymer solution. Becauseturbulent flow of the two-phase fluid allows it to absorb heat much morerapidly than a laminar flow of a viscous polymer solution, it ispossible to heat the two-phase fluid to the proper devolatilizationtemperature in approximately one-third the time required to heat asingle phase of polymer/monomer liquid. It is important that thepreheater be designed with a relatively large number of tubes of a largeenough diameter so that the vacuum may be pulled back through the tubesand so that there is formed the two-phase, preferably foamed or frothedfluid therein. Preferably, the length and diameter of the tubes 42 arechosen such as to permit this two-phase flow.

In addition to decreasing the residence time of the polymer in the heatexchanger or preheater 41, the early boiling of the volatiles to formthe two-phase fluid provides the added advantage that the acrylonitrileis now in the vapor phase. As discussed above, acrylonitrile monomerdoes not readily polymerize in the vapor phase. This further reduces thepossibility of the formation of SAN having non-uniform composition (andthus an off-color copolymer).

A third advantage is realized when the two-phase fluid is introducedinto devolatilizer 45. In the prior art, when a pressurized liquidpolymer solution of the appropriate temperature for devolatilization wasintroduced into the devolatilizer, the monomer would boil at theinjection nozzle. This boiling would reduce the temperature of thepolymer by as much as 25° C. as the volatiles absorbed their latent heatof vaporization. This phenomenon would result in a colder polymer whichis thus more difficult to pump and, therefore, requires more energy. Inaddition, the temperature of the melt fluctuates at the die by as muchas about 25° C., causing problems in pelletizing, e.g., surging andimproper temperature control leading to pellets having poor properties.

By contrast with the prior art, in the present system the seconddevolatilizer functions mainly as a phase separation chamber. A constanttemperature polymer is produced, and there is little cooling as thepolymer drops down through the second devolatilizer. The devolatilizedpolymer is withdrawn through line 34 and has a purity of about 99.80%.The remaining 0.20% consists essentially of volatiles. If an extremelyhigh purity SAN copolymer is desired, another devolatilizer of the sametype as devolatilizer 45 may be provided downstream of devolatilizer 45.The overhead product of diluent, acrylonitrile monomer and styrenemonomer is withdrawn through line 49 for further processing and/orinjection into reactor 1.

As indicated above, a vacuum is pulled in the second devolatilizer 45and in part also in the first stage devolatilizer 35, in order towithdraw diluent and unpolymerized monomer which are subsequentlycondensed and preferably returned to the polymerization reactor. Anyconventional source of vacuum, designated by reference numeral 51, maybe employed in the system according to the present invention. Forexample, a conventional steam jet vaucum source may be employed.However, it is also possible, and it is preferred according to theinvention, to employ a relatively simple low pressure system which usesrotary blowers as the source of negative pressure. The system accordingto the invention does not require refrigeration, thereby eliminatingcostly equipment and a significant operating expense factor. It is alsopossible to employ a liquid sealed vacuum pump; however, this produces alarge quantity of contaminated water, an environmental problem which isalso obviated by the preferred system according to the invention.

The volatile components from the devolatilizing system are drawn by thevacuum from vacuum source 51 into recycle condenser 53, which is cooledby cooling water. Condensed monomers and diluent pass from recyclecondenser 53 to recycle surge drum 55, from which they are pumped byrecycle pump 57 via line 59 back to reactor 1. These recycled componentsare preferably introduced upstream of preheater 8.

In FIG. 3 is illustrated a preferred recycle condenser system accordingto the invention. Recycle condenser 53 is fed with cooling water vialine 61, and the cooling water leaves the condenser vai line 63. Thecondenser 53 receives the volatilized constituents (diluent andmonomers) from the devolatilizer system, and after the volatiles arecondensed, they are transported in liquid form to recycle surge drum 55via line 65. Recycle surge drum is provided with a trap 56 for removalof any water in the system.

Condenser 53 is connected to the vacuum supply via line 67, and thepressure is equalized between the condenser and surge drum 55 byproviding interconnection through line 69. The vacuum in the recyclecondenser system is limited by the hydrocarbon vapor pressure in surgedrum 55, and therefore, it is desirable to maintain the temperature aslow as possible. Typically, the temperature is maintained between about27° and 38° C.

Liquid monomer and diluent are withdrawn from surge drum 55 and pumpedby recycle pump 57 through line 59 back to the polymerization reactor,as described above. Provision is also made to selectively recycle aportion of the condensate through line 58 back to the surge drum 55and/or a portion thereof through line 60 back to the condenser 53. Inthe latter case, the recycled liquid is advantageously injected througha spray nozzle into the incoming stream of volatiles.

Condenser 53 is designed to produce a pressure drop thereacross of lessthan about 2 mm Hg, and it is sized such that there is less than about2° C. approach between the cooling water and the hydrocarbon leaving it.Typically, the cooling water entering is at a temperature between about16° and 29° C., and the water leaving is between about 20° and 32° C.The pressure in the condenser system is typically between about 0.2 and0.3 psia.

The following example will further illustrate the present invention ingreater detail; however, it is to be understood that the scope of thisinvention is not to be limited by this example:

EXAMPLE 1

In a vertical stirred tank reactor having a volume of 6,000 gals. amixture of styrene and acrylonitrile is copolymerized. A continuousstream of reaction mixture is withdrawn from the bottom of the reactor,passed through an external loop which incorporates a static mixing zoneand is reintroduced into the top of the reactor at a rate sufficient toturn over the reaction mixture about 1.2 times per hour. The mixture isconveyed by a positive displacement type pump which is operated at aconstant speed. An uninterrupted feed stream of liquid monomer composedof about 70 weight percent styrene and 30 weight percent acrylonitrileis charged into the circulating stream of reaction mixture just prior tothe point at which it enters the static mixing zone at a rate of about4,000 pounds per hour. An additional charge of recycled monomerscomposed of about 28 weight percent styrene, 34 weight percentacrylonitrile, and 38 weight percent diluent is continuously fed at arate of about 4,400 lbs. per hour into the external recirculation loopat the same point as the fresh feed stream. The combination of therecycle and feed monomers is thoroughly mixed with the reaction mixturevia the static mixing zone. This mixture is then passed back into thereactor.

A thoroughly homogeneous composition is maintained in the reaction zoneby a combination of the external recirculation loop and agitation withan agitator suspended from the top of the reactor. The agitator includesboth anchor and turbine-type blades turning at a speed of 30 rpm. Thiscombined agitation is sufficient to insure a maximum temperaturegradient of ±1° C.

The polymerization reactor is jacketed and a heat transfer oil iscirculated through the jacket. The temperature of the oil is maintainedat a constant temperature of about 144° C. by an automatic control.

Upon reaching steady state, the liquid contents of the reactor are heldat a constant temperature of about 144° C. and at a pressure of about 40psia. A constant liquid volume is maintained in the reactor bycontinuously withdrawing a portion of the partially polymerized reactionmixture from the bottom at a rate equal to that of the combined feed andrecycle streams entering the reactor. This stream is composed of about48 weight percent styrene/acrylonitrile copolymer, 14.5 weight percentstyrene monomer, 17.5 weight percent acrylonitrile monomer and 20 weightpercent diluent.

The copolymer contains about 29 weight percent acrylonitrile and 71weight percent styrene and is of a homogeneous composition. There is anabsence of color, with the polymer being essentially clear.

The heat of polymerization is removed by boiling monomers from thereaction zone at a controlled rate. A vertical heat exchanger of theshell and tube type located adjacent to the reactor is used to condensethe hot vapors, the vapors being condensed on the inside of the tubesand the shell being circulated with cooling water.

A liquid level is maintained in the tubes of the condenser, with themonomers being returned immediately to the reactor at the same point ofentry as the feed and recycle streams. The rate of cooling is controlledby adjusting the level of liquid monomer in the tubes of the condenser,thus adjusting the area available for condensation. A sidestream ofcooled monomers is sprayed into the top of the condenser above the pointat which the vapors from the reactor enter. The vapor stream, composedof about 8 weight percent styrene, 78 weight percent acrylonitrile, and14 weight percent diluent enters the condenser at about 144° C. A volumeof condensate from the condenser is stored in a collection tankpositioned between the condenser and the condensate recycle pump. Thisvolume amounts to approximately 500 gals.

The copolymer/monomer mixture exiting the reactor is forwarded to thedevolatilization section by a variable speed gear pump. This pumpmaintains a constant level in the reaction zone. This mixture is passedthrough a jacketed pipe circulated with heat transfer oil which is runat the same temperature as that in the reactor jacket and enters thefirst of the two devolatilization vessels. The liquid mixture is at atemperature of about 144° C. The first vessel is maintained at apressure of about 29 psia by an automatic pressure control. Vaporscomposed of about 19 weight percent styrene, 51 weight percentacrylonitrile and 30 weight percent diluent are boiled out of thecopolymer mixture as it enters the vessel at a rate of about 2,200pounds per hour. The resulting polymer mixture composed of about 13weight percent styrene, 5.5 weight percent acrylonitrile, 16.5 weightpercent diluent, and about 65 weight percent styrene/acrylonitrilecopolymer at about 77° C. falls to the bottom of the vessel and isforwarded out by a pump. The ratio of styrene and acrylonitrile monomersis approximately the same as in the feed ratio to the polymerizationreactor.

From the bottom the copolymer solution enters the tube side of avertical shell and tube type heat exchanger, in which it is heated to atemperature of about 230° C. by heating oil circulated on the shellside. The design of this preheater is such that a vaporization of themonomer occurs within the tubes and a foaming mixture of monomer/polymeris formed. The flow characteristics of this foamed mixture is such thatit allows for much better heat transfer than with a conventionalexchanger of this type, allowing for a much smaller heating zone. Thetubes are sufficiently large to allow the vacuum from the second stageof the devolatilizer to extend into them and to permit foaming of themonomer/polymer mixture therein. The length of each tube is 12 feet andthe diameter is 11/4 inches.

The two-phase mixture exits the preheater at about 233° C. and entersthe second devolatilization vessel through a distribution nozzle whereit is separated into small strands. The vessel is operated at a pressureof about 0.29 psia and at 230° C. A mixture of about 37 weight percentstyrene, 16 weight percent acrylonitrile, and 47 weight percent diluentis taken off as a vapor, is condensed and is recycled and returned tothe reaction vessel. The resulting devolatilized copolymer composed ofabout 0.15 weight percent styrene, 0.01 weight percent acrylonitrile,0.05 weight percent diluent, and 99.79 weight percent copolymer falls tothe bottom of the vessel and is passed through a pump to the pelletizingdies.

EXAMPLES 2-6

The procedure of Example 1 is repeated with various changes inparameters. The parameters for these additional Examples, as well asExample 1, are summarized in the following table.

    __________________________________________________________________________               Re-            FEED             COPOLYMER                                POLY ACTOR                                                                              AGI- JACKET   %       %        %           MELT               EXAM- TEMP.                                                                              PRESS.                                                                             TATOR                                                                              TEMP.                                                                              RATE                                                                              STY-    POLY-                                                                              RATE                                                                              STY-        FLOW               PLE No.                                                                             °C.                                                                         (psia)                                                                             RPM  °C.                                                                         lbs/hr                                                                            RENE                                                                              ACN MER  lb/hr                                                                             RENE                                                                              ACN MW  G/10'              __________________________________________________________________________    1     144  40   30   144  4,000                                                                             70  30  50   4,000                                                                             74  26  120,000                                                                           14.0               2     155  50   26   155  4,000                                                                             70  30  50   4,000                                                                             74  26  105,000                                                                           27.0               3     130  30   26   130  4,000                                                                             70  30  50   4,000                                                                             74  26  155,000                                                                           3.0                4     144  40   30   144    200                                                                             70  30  50     200                                                                             74  26  120,000                                                                           14.0               5     130  30   26   130  4,000                                                                             60  40  50   4,000                                                                             67  33  150,000                                                                           3.0                6     155  50   28   155  20,000                                                                            70  30  50   20,000                                                                            74  26  105,000                                                                           25.0               __________________________________________________________________________

What is claimed is:
 1. A process for the continuous mass polymerizationof styrenic and alkenylnitrile monomers to produce astyrenic/alkenylnitrile copolymer, comprising the steps of:(a)continuously introducing a feed comprising a predetermined ratio ofstyrenic and alkenylnitrile monomers into a reation vessel to produce areation mixture; (b) subjecting the reaction mixture containing styrenicand alkenylnitrile monomers to conditions of temperature and pressureunder which said monomers copolymerize to produce a liquid containingstyrenic/alkenylnitrile copolymer and styrenic and alkenylnitrilemonomers; (c) subjecting the reaction mixture to agitation sufficient tomaintain a substantially uniform composition distribution and asubstantially uniform temperature distribution throughout the reactionmixture; (d) continuously withdrawing at least part of a liquidcontaining styrenic/alkenylnitrile copolymer and styrenic andalkenylnitrile monomers from the reaction vessel; and (e) cooling thereaction mixture by withdrawing vapor phase containing vaporizedstyrenic and alkenylnitrile monomer from the reaction vessel to acondenser, condensing the vaporized styrenic and alkenylnitrilewithdrawn from the reaction vessel by contact with a cooled condensingsurface to produce a condensed monomer-containing liquid and returningat least a part of the condensed monomer-containing liquid to thereaction vessel, wherein said cooling step further comprises controllingthe amount of vaporized monomers withdrawn from the reaction vessel byvarying the amount of cooled surface area available for contact with thewithdrawn vaporized monomers in response to the temperature in thereaction vessel, whereby the amount of vaporized monomers condensing iscontrolled.
 2. A process as defined in claim 1, wherein said coolingstep further comprises collecting a volume of the condensedmonomer-containing liquid and holding same for selective introductioninto the reaction vessel to produce an instantaneous cooling effect. 3.The process of claim 1 wherein varying the amount of cooled surface areaavailable for contact with the withdrawn vaporized monomers is achievedby varying the amount of condensed monomers retained in the condenser.4. A process as defined in claim 1, wherein said cooling step furthercomprises the step of recycling a portion of the condensedmonomer-containing liquid to the condensing step and quenching thevaporized monomer withdrawn from the reaction vessel by passing thewithdrawn vaporized monomer through a stream of the recycled condensedmonomer-containing liquid.
 5. A process as defined in claim 4, whereinsaid stream of recycled condensed monomer-containing liquid comprises acurtain-like distributed stream.
 6. A process as defined in claim 4,wherein said cooling step further comprises the step of impinging saidrecycled condensed monomer-containing liquid against the cooled surfacearea to prevent accumulation of polymerized monomer on said surface. 7.A process as defined in claim 1, wherein said cooling step furthercomprises contacting the outside of the walls of the reaction vesselwith a liquid cooling medium.
 8. A process as defined in claim 1,further comprising the step of purging non-condensible components fromthe vapor phase withdrawn from the polymerization vessel.
 9. A processas defined in claim 8, wherein said step of purging non-condensiblescomprises collecting said non-condensibles in a region located above thecooled condensing surface and selectively withdrawing thenon-condensibles from the collection region.
 10. A process as defined inclaim 9, wherein said step of selectively withdrawing non-condensiblescomprises detecting the interface between the collected non-condensiblesand the condensible vapors of the withdrawn vapor phase and venting thenon-condensibles in response to movement of said interface to apreselected position indicating collection of a preselected volume ofnon-condensibles.
 11. A process as defined in claim 3, wherein saidcooling step further comprises directly venting said vapor phase fromthe reaction vessel through an unobstructed line to the inlet of acondenser positioned above the reaction vessel.
 12. A process as definedin claim 1, further comprising the step of removing volatile componentsfrom the liquid containing styrenic/alkenylnitrile copolymer, styrenicmonomer and alkenylnitrile monomer withdrawn from the reaction vessel.13. A process as defined in claim 12, wherein said devolatilizing stepcomprises:(a) in a first stage, adjusting the ratio of styrenic toalkenylnitrile monomer in the withdrawn liquid to a ratio which isselected to provide, upon polymerization, a styrenic/alkenylnitrilecopolymer having approximately the same ratio of styrenic:alkenylnitrilemonomer units as the styrenic/alkenylnitrile copolymer withdrawn fromthe reaction vessel; and (b) in a second stage, heating the withdrawnliquid to a temperature sufficient to volatilize substantially all ofthe volatile components therein and separating the volatilizedcomponents from the non-volatilized copolymer.
 14. A process as definedin claim 13, wherein said adjusting step comprises withdrawing asufficient amount of alkenylnitrile monomer.
 15. A process as defined inclaim 14, wherein said withdrawing comprises maintaining the pressure ofthe first stage at a level sufficient to provide a monomer-containingvapor phase and selectively venting alkenylnitrile monomer-rich vapors.16. A process as defined in claim 13, wherein said heating step in saidsecond stage comprises passing the withdrawn liquid through a heatexchange apparatus and reducing the pressure in the heat exchangeapparatus by an amount sufficient to cause substantial volatilization ofthe volatile components in the liquid sufficient to cause a froth-typetwo-phase flow through the heat exchange apparatus.
 17. A process asdefined in claim 16, wherein said separating of volatilized componentsin said second stage comprises passing the two-phasecopolymer/volatilized components stream from the heat exchange apparatusinto a separation vessel and reducing the pressure in the separationvessel to a level sufficient to maintain the volatilized components invapor form.
 18. A process as defined in claim 13, wherein said processfurther includes the steps of withdrawing the volatilized componentsfrom the separation vessel in the second stage, condensing the withdrawnvolatilized components to produce a liquid condensate of recoveredmonomers, and recycling the recovered monomer liquid to the reactionvessel as part of the monomer feed.
 19. A process as defined in claim 1,wherein the feed contains from about 10 to 60% by weight ofacrylonitrile and from 40 to 90% by weight of styrene, based upon thetotal monomer content.
 20. A process as defined in claim 1, wherein thefeed further comprises from about 2 to 50% by weight of an inert diluentfor the monomers.
 21. A process as defined in claim 20, wherein saidinert diluent comprises ethylbenzene, butylbenzene, benzene, toluene,xylene or cumene.
 22. A process as defined in claim 1, wherein thecopolymerization conditions are such as to produce a copolymerizationrate of up to about 40% per hour and a total conversion of copolymer ofup to about 90% based on the monomer feed.
 23. A process as defined inclaim 1, further comprising the step of maintaining not above anequilibrium concentration of water in the reaction vessel.